Process to produce ethylene and hydrogen

Abstract

The present disclosure provides a process to produce ethylene and hydrogen comprising: providing a first reactant mixture (15) comprising methane, oxygen; performing an Oxidative Coupling of Methane (OCM) reaction on the first reactant mixture (15) to produce a first product stream (17) comprising carbon monoxide, ethylene and hydrogen; providing a second reactant mixture (19); performing a steam cracking reaction on the second reactant mixture (19) to produce a second product stream 21 comprising ethylene and hydrogen; separating at least a part of the carbon monoxide; and performing a water gas shift reaction on the CO-rich stream 71 to produce an H2-rich stream 83; wherein the OCM reaction and the steam cracking reaction are performed simultaneously in separate chambers of a single reactor being a reactor 1, the reactor 1 is configured to transfer the heat generated by the OCM reaction to the steam cracking reaction.

Description

[0001]PROCESS TO PRODUCE ETHYLENE AND HYDROGENTechnological fieldThe present disclosure relates to a process to produce ethylene and hydrogen involving awater gas shift reaction and a coupling between a steam cracking reaction and an oxidativecoupling of methane (OCM) reaction so that the heat generated by the exothermic OCMreaction is transferred to the endothermic steam cracking reaction. Technical backgroundClimate change and ongoing energy transition make it mandatory to replace fossil carbon-based fuels in chemical production and recycled processes with a more environmentally friendly decarbonized source of energy. Transforming natural gas into valuable chemicalsrequires elevated temperatures, often higher than 800°C and even up to 1,000°C, and areoften endothermic. The energy needed is, therefore, high and not often environmentally friendly, as it is demonstrated by the common use of fired heated reactors. An example of an endothermic reaction that is commonly performed by fired heated reactors is the steam cracking reaction that allows the production of ethylene.Steam cracking is a high-temperature, low-pressure chemical process that is used to breakdown larger hydrocarbon molecules into smaller, more valuable molecules. The steam cracking reaction is well-known to the person skilled in the art and is described for example in US11046893B2 and in WO2009025640A1. The process is a key step in the production of various petrochemicals, including ethylene and propylene, which are used to make plastics, solvents, and other chemical products. In steam cracking, a feedstock, typically a liquid or gaseous hydrocarbon such as naphtha, propane, or ethane, is mixed with steam and heated to temperatures ranging from 750°C to 950°C. The feedstock is then rapidly cracked or broken down into smaller molecules, such as ethylene and propylene, through a series of chemical reactions involving the breaking of carbon-carbon bonds. The exact reactions that take place in steam cracking depend on the feedstock and reaction conditions used. Generally, steam cracking involves the following types of reactions: -Pyrolysis: The feedstock is broken down into smaller molecules through the breakingof carbon-carbon bonds. This results in the formation of smaller hydrocarbon molecules, such as ethylene and propylene.- Polymerization: The smaller hydrocarbon molecules formed in pyrolysis can undergofurther reactions to form larger molecules through the formation of carbon-carbon bonds. This can result in the formation of undesirable by-products or coke.- Cracking: The larger hydrocarbon molecules that form in polymerization can bebroken down into smaller, more desirable molecules through further cracking reactions. The products of steam cracking are a mixture of hydrocarbons, including ethylene, propylene, butenes, and other byproducts. The product mixture is typically separated and purified to isolate the desired product, such as ethylene, which can then be used as a feedstock in various chemical reactions. Steam cracking is a very energetic process for the production of ethylene and H2. Nowadays, the heat needed for the reaction is supplied through the combustion of methane, responsiblefor releasing large amounts of CO2. A solution to avoid such a release is to be found.To contribute to the replacement of the use of fossil carbon-based fuels heating devicesEP3945066 is disclosing a process and an installation to perform a steam cracking reaction,wherein the heat is electrically generated. The use of green electricity is contemplated, i.e., electricity coming from solar energy and / or wind energy. The process disclosed comprises the steps of providing a fluidized bed reactor comprising at least two electrodes; and a bed comprising particles, wherein the particles are put in a fluidized state by passing upwardly through the said bed a fluid stream, to obtain a fluidized bed; heating the fluidized bed to a temperature ranging from 500°C to 1200°C to conduct the endothermic chemical reaction; wherein at least 10 wt.% of the particles based on the total weight of the particles of the bedare electrically conductive particles and have a resistivity ranging from 0.001 Ohm.cm to 500Ohm.cm at 800°C and in that the step of heating the fluidized bed is performed by passing an electric current through the fluidized bed. This solution was found very interesting. Nevertheless, other ways of producing ethylene involving steam cracking reaction without using fired heated reactors need to be found. Inparticular, processes that could be less energy-demanding.Also, there is a need for enhancing hydrogen production.CN 110041 156 A relates to the technical field of methane conversion, in particular to anintegrated process for directly converting methane to ethylene, wherein methane oxidativecoupling reaction and alkane cracking reaction are simultaneously carried out in different cavities of a coupling reactor, methane and oxygen-containing raw gas entering a first cavity is subjected to the methane oxidative coupling reaction under the action of a first catalyst torelease heat, the heat is transferred to a second cavity through an interlayer wall, an ethane and C2+ mixture entering the second cavity are subjected to the alkane cracking reaction under the action of a second catalyst to absorb the heat, and the coupling reactor is a jacketed packed bed type formed by a tube side and a shell side or a tubular tube and shell side heat exchange type coupling reactor.CN 113 800 995 A discloses a method and a system for coupling propane catalyticdehydrogenation reaction and methane oxidative coupling reaction. The method is carried out in a fixed bed sleeve reactor which contains at least two layers of sleeves. The methane oxidative coupling reaction is carried out in at least one of the sleeves, and the propane catalytic dehydrogenation reaction is carried out in at least another adjacent sleeve.US 5 254 781 A discloses an olefins process that combines cracking of a hydrocarbonfeedstock with the coupling of methane using an oxygen-affording gas such that the heat evolved in the exothermic methane coupling reaction is effectively transferred to theendothermic cracking process in a manner that does not allow the non-hydrocarbon productsin the effluent of the methane coupling reaction to mix with the effluent of the cracking process. By combining the cryogenic requirements of the two processes, the refrigerationused in air liquefaction to separate oxygen for the methane coupling process can providerefrigerant nitrogen to liquify one or more of methane, ethylene and propylene which can be used in the cryogenic separation of the C2+ hydrocarbon products in the cracking process effluent.EP 4082996 A1 provides for a process for a non-catalytic oxidative coupling of methanereaction comprising a step of providing a counter-current shell-tube reactor comprising atleast two tubes defining a tubular part and a shell part surrounding the tubular part and at least one inlet to feed a gaseous feed stream and at least one outlet to discharge a product stream; a step of providing a gaseous feed stream comprising a gas mixture of methane and oxygen in a defined molar ratio and preheated to a defined operating inlet temperature; a step of feeding the gaseous feed stream at least in the tubular part of the counter-current first reactor section-tube reactor and a step of recovering a product stream. The present disclosure aims to provide a large-scale solution to produce ethylene and hydrogen that is suitable for a large-scale application in the industry, such as the chemical industry. The present disclosure aims to provide an ethylene and hydrogen production processinvolving a steam cracking reaction with enhanced hydrogen production, wherein the processcontributes to the replacement of the use of fossil carbon-based fuel heating devices in steamcracking reactors. The present invention provides a solution to conduct endothermic steamcracking of hydrocarbons into ethylene with reduced energy supply and / or that reduces oravoids the use of fired heated reactors. Summary of the disclosureIt has now been found that is possible to reply to one or more of the needs encountered inthe prior art by producing ethylene with a thermal coupling between a steam crackingreaction and an oxidative coupling of methane in a single reactor so that the heat generatedby the exothermic OCM reaction is transferred to the endothermic steam cracking reaction; wherein the process integrates a water-gas shift.In a first aspect, the invention provides a process to produce ethylene and hydrogenremarkable in that it comprises:- providing a first reactant mixture comprising methane, oxygen, and optional carbondioxide; -performing an Oxidative Coupling of Methane (OCM) reaction on the first reactantmixture to produce a first product stream comprising carbon monoxide, ethylene, andhydrogen, wherein the OCM reaction is performed in the absence of a catalyst at a temperature ranging from 750°C to 1200°C and a pressure ranging from 0.1 to 5.0 MPa: -providing a second reactant mixture comprising hydrocarbons containing two or morecarbon atoms in their molecular structure; -performing a steam cracking reaction on the second reactant mixture in absence of acatalyst to produce a second product stream comprising ethylene and hydrogen;- separating at least a part of the carbon monoxide contained in, or originating from, thefirst product stream to produce a CO-rich stream; and -performing a water gas shift reaction on the CO-rich stream to produce a H2-rich streamwherein the OCM reaction and the steam cracking reaction are performed simultaneously inseparate sections of a single reactor having two reactor sections, the reactor beingconfigured to transfer the heat generated by the OCM reaction in the first reactor section tothe second reactor section for the steam cracking reaction.As it is understood from the above definition, the present disclosure comprises a heatintegration process for a greener production of ethylene together with an improved hydrogenproduction, wherein a single reactor (such as a shell and tube heat exchange-like reactor) isused so the exothermic OCM reaction provides the heat for the steam cracking reaction(endothermic), and wherein hydrogen either directly by the steam cracking reaction or indirectly by the OCM reaction by performing a water-gas shift reaction on the CO produced by the OCM reaction. Indeed, oxidative coupling of methane (OCM) is an alternative route for ethylene andpropylene production that is currently obtained through steam cracking of higherhydrocarbons. However, apart from an alternative, OCM can also be an additional route tosteam cracking, keeping in mind that methane (CH4) is a by-product of steam cracking and istypically burnt as fuel for heating the same steam crackers. OCM is an exothermic process,and one of the main challenges consists of controlling the heat released during the reaction,to avoid overoxidation and thus high amounts of CO2at the expense of the production of ethane and ethylene.The process is remarkable in that the endothermic steam cracking acts as a quencher to theOCM reaction, avoiding temperature runaway in the OCM reactor. Despite the heatefficiency, the process of the disclosure also allows for the reduction of CO2 emissionsrelated to steam cracking, once part of the steam cracking of hydrocarbons, containing twoor more carbon atoms in their molecular structure are replaced by the OCM reactor,integrated with a steam cracking section in one single reactor. Furthermore, besidesproviding a heat source, OCM also converts methane into ethylene and syngas whichcomprises CO. The process of the disclosure takes advantage of the CO presence in thefirst product stream to recover it as a CO-rich stream to produce additional hydrogen via thewatergas shift reaction. Attempts to use the heat generated by the OCM reaction have already been proposed. Forexample, WO2015081122A2 discloses a process based on a two-section reactor, whereinthe first section methane is converted into ethane and ethylene through OCM, and in the second section, the heat from OCM is used for ethane cracking (co-produced in the first section and imported ethane). This is different from the present disclosure wherein the reactions are performed in parallel in different sections of a heat exchange-like reactor, andthe OCM effluent is not sent to the ethane cracking section. In the installation, the crackinginlet stream is independent of the OCM reactor. This configuration was found to be particularly advantageous for non-catalytic OCM wherein the ethylene (C2= olefin) / ethane (C2 paraffin) is elevated, reducing the need for further cracking.WO2016094482A1 discloses a method of producing ethylene and synthesis gas bycombining OCM and dry reforming of methane reactions. This document discloses the two reactions being performed in the same reactor using one or more catalysts active for both OCM and dry reforming.WO2020142594A1 discloses the combination of OCM and catalytic partial oxidation (CPO)of methane. It is a two-zone reaction, where the OCM reaction zone is followed by CPO to produce mainly syngas and C2 products.One or more of the following can be used to further describe the process according to thedisclosure. Different types of heat exchange-like reactors can be used in the process of the first aspect,i.e., for the integration of both OCM and steam cracking. However, they must all allow theheat exchange between a first reactor section comprising one or more chambers whereinthe OCM reaction occurs and a second reactor section comprising one or more chambers wherein the steam cracking reaction occurs. In a preferred embodiment, the reactor is ashell-tube reactor comprising a first reactor section being a shell and a second reactorsection comprising one or more inner tubes; and the OCM reaction takes place in the shell and the steam cracking reaction takes place in the one or more inner tubes. The first reactorsection (i.e. the shell) can form a unique chamber or comprise one or more chambers;preferably, the first reactor section (i.e. the shell) forms a unique chamber.Shell-tube reactors are well-known to the person skilled in the art and can have severalconfigurations.In a preferred configuration, the shell-tube reactor is a co-current reactor comprising a firstreactor section being a shell and a second reactor section comprising one or more inner tubes placed inside the shell. The one or more inner tubes can be straight tubes or U-shaped tubes. OCM takes place in the shell and steam cracking in the one or more inner tubes. Theuse of preheaters is preferred wherein the temperature of the one or more preheaters shouldbe between 350-700°C.Thus, in such an embodiment, the shell-tube reactor is a co-current shell-tube reactorcomprising a first reactor section being the shell and a second reactor section being one ormore inner tubes placed inside the shell; wherein the OCM reaction takes place in the shelland the steam cracking reaction takes place in the one or more inner tubes.With preference, the first reactant mixture and / or the second reactant mixture is pre-heatedbefore entering the reactor. For example, the first reactant mixture and / or the secondreactant mixture is pre-heated to a temperature ranging from 350 to 700°C.Whatever the configuration used, the following can be used to better define the steamcracking reaction and the OCM reaction.Regarding the OCM reactionThe OCM reaction is performed through a non-catalytic (thermal) approach. In other terms, theOCM reaction is performed in the absence of a catalyst.According to the disclosure, the OCM reaction is performed in the absence of a catalyst, at atemperature ranging from 750°C to 1200°C and a pressure ranging from 0.1 to 5.0 MPa. Thelight-off temperatures are preferably above 750 °C, more preferably above 850 °C. Preferably,the reactor volume to methane flow ratio (V / F (CH4)) is between 0.002 and 1 m3.s.mol-1. Withpreference, the pressure ranges from 0.1 to 2.0 MPa. Feed to the OCM reactor (shell side) is a mixture of CH4 / O2or CH4 / air. It is preferred that thefirst reactant mixture shows an operating molar ratio CH4:O2 ranging from 2.5 to 20.0,preferably, 2.5 to 10.0 or 5.0 to 10.0.Regarding the steam cracking reactionThe steam cracking reaction is also done in the absence of a catalyst. The steam crackingreaction is performed on a second reactant mixture comprising hydrocarbons containing twoor more carbon atoms in their molecular structure, wherein the second reactant mixturepreferably comprises one or more selected from ethane, propane, butane, liquefiedpetroleum gas, naphtha, and diesel; with preference, the second reactant mixture comprises one or more selected from ethane, propane, and naphtha; more preferably, the second reactant mixture comprises ethane and / or propane.For example, the steam cracking reaction is performed at a temperature ranging from 750 to900°C and / or at a pressure ranging from 0.1 to 0.5 MPa.For example, the residence time of the second reactant mixture in the reactor ranges from0.10 to 0.30 seconds.The step of separating at least a part of the carbon monoxideIn an embodiment, the step of separating at least a part of the carbon monoxide containedin, or originating from, the first product stream to produce a CO-rich stream also produces ahydrogen-rich stream. In an embodiment, the step of separating at least a part of the carbon monoxide containedin, or originating from, the first product stream to produce a CO-rich stream is performed bycryogenic distillation, or by pressure swing adsorption (PSA) or by COSORB process; preferably by PSA adsorption. The water-gas shift reaction In an embodiment, the process comprises transferring a part of the heat generated by thewater-gas shift reaction to pre-heat the first reactant mixture and / or the second reactantmixture before entering the reactor; preferably, using a heat exchanger. Thus, the process comprises transferring a part of the heat generated by the water-gas shiftreaction to pre-heat the first reactant mixture before it enters the first reaction section; and / orthe process comprises transferring a part of the heat generated by the water-gas shiftreaction to pre-heat the second reactant mixture before it enters the second reactor section.Other possible stepsIn an embodiment, the pressure of the first product stream is higher than the pressure of thesecond product stream, and the process comprises a compression step wherein the secondproduct stream is compressed to at least the pressure of the first product stream andsubsequently mixed with the first product stream to form a mixed stream before the step ofseparating at least a part of the carbon monoxide contained in, or originating from, the firstproduct stream to produce a CO-rich stream. In an embodiment, before the step of separating at least a part of the carbon monoxidecontained in, or originating from, the first product stream to produce a CO-rich stream; atleast one of the first product stream, the second product stream, or the compressed gasstream (if applicable) is submitted to a purification step to remove acid gas; with preferencethe purification step comprises a sub-step of amine washing followed by a sub-step of caustic washing. For example, before the step of separating at least a part of the carbon monoxide containedin, or originating from, the first product stream to produce a CO-rich stream; the first productstream is submitted to a purification step to remove acid gas; with preference the purificationstep comprises a sub-step of amine washing followed by a sub-step of caustic washing.With preference, the purified stream and the compressed gas stream are mixed to form amixed stream before the step of separating at least a part of the carbon monoxide originatingfrom the first product stream to produce a CO-rich stream.In an embodiment, before the step of separating at least a part of the carbon monoxidecontained in, or originating from, the first product stream to produce a CO-rich stream; atleast one stream selected from the first product stream, the second product stream and a mixed stream is submitted to a step of hydrocarbon recovery to recover a stream containing C2+ products; with preference, the step of hydrocarbon recovery is performed by cryogenic separation.The terms “C2+ products” refer to hydrocarbons that contain two or more carbon atoms intheir molecular structure.According to a second aspect, the disclosure provides for an installation to produce theprocess according to the first aspect wherein the installation comprises, fluidically connected in the following order: -a reactor having two reactor sections and configured to transfer the heat generatedin a reactor section to another reactor section; -an optional compression unit;- an optional purification unit;- an optional drying unit;- a hydrocarbon separation unit;- a CO removal unit;- a water-gas shift unit; and- an optional hydrogen separation unit.With preference, the reactor is a shell-tube reactor comprising a first reactor section being ashell and a second reactor section comprising one or more inner tubes. Description of the figures -Figure 1 illustrates an example of a configuration of a reactor having two reactorsections to perform the coupling OCM-SC in a single reactor.- Figure 2 illustrates another example of a configuration of a reactor having two reactorsections to perform the coupling OCM-SC in a single reactor.- Figure 3 illustrates the configuration of the reactor having two reactor sections usedin the first example. -Figure 4 illustrates the temperature profile along the reactor used in the first example.- Figure 5 illustrates an embodiment of the process according to the disclosure.- Figure 6 illustrates an installation to perform the process according to the disclosure.Detailed descriptionFor the disclosure, the following definitions are given: The terms "comprising", "comprises" and "comprised of" as used herein are synonymous with "including", "includes" or "containing", "contains", and are inclusive or open-ended and do notexclude additional, non-recited members, elements, or process steps. The terms "comprising”,"comprises" and "comprised of" also include the term “consisting of”. The recitation of numerical ranges by endpoints includes all integer numbers and, whereappropriate, fractions subsumed within that range (e.g., 1 to 5 can include 1, 2, 3, 4, 5 when referring to, for example, a number of elements, and can also include 1.5, 2, 2.75 and 3.80, when referring to, for example, measurements). The recitation of endpoints also includes therecited endpoint values themselves (e.g., from 1.0 to 5.0 includes both 1.0 and 5.0). Anynumerical range recited herein is intended to include all sub-ranges subsumed therein.The particular features, structures, characteristics, or embodiments may be combined in anysuitable manner, as would be apparent to a person skilled in the art from this disclosure, inone or more embodiments. Reference is made to Figure 6 unless indicated otherwise.The present disclosure provides a process to produce ethylene and hydrogen remarkable inthat it comprises:- providing a first reactant mixture 15 comprising methane, oxygen, and optionalcarbon dioxide; -performing an Oxidative Coupling of Methane (OCM) reaction on the first reactantmixture 15 to produce a first product stream 17 comprising carbon monoxide, ethylene, and hydrogen; wherein the OCM reaction is performed in the absence of acatalyst, at a temperature ranging from 750°C to 1200°C and a pressure ranging from 0.1 to 5.0 MPa; -providing a second reactant mixture 19 comprising hydrocarbons containing two ormore carbon atoms in their molecular structure; -performing a steam cracking reaction on the second reactant mixture 19 in theabsence of a catalyst to produce a second product stream 21 comprising ethyleneand hydrogen; -separating at least a part of the carbon monoxide contained in, or originating from,the first product stream to produce a CO-rich stream 71; and- performing a water gas shift reaction on at least a part of the carbon monoxidecontained in, or originating from, the first product stream to produce an H2-rich stream83; wherein the OCM reaction and the steam cracking reaction are performed simultaneously inseparate sections of a single reactor, having two reactor sections (3, 5), the reactor 1 beingconfigured to transfer the heat generated by the OCM reaction in the first reactor section tothe second reactor section for the steam cracking reaction;with preference, the OCM reaction and the steam cracking reaction are performed in areactor 1 having two reactor sections (3, 5) which is a shell-tube reactor with a first reactorsection 3 being a shell and a second reactor section 5 comprising one or more inner tubes;wherein the OCM reaction takes place in the shell and the steam cracking (SC) reactiontakes place in the one or more inner tubes so the heat generated by the OCM reaction istransferred to the steam cracking reaction.Heat integration must be accessed in detail to guarantee the best optimization for the process, assuring control of exothermicity on OCM reaction and enough transfer of caloriesto steam cracking. Mass flow rates, residence time, and reactor sizing for both OCM andsteam cracking units must be selected to ensure the conditions of sufficient heat generationby the OCM reaction, based on the heat demand from the steam cracking reaction.The disclosure also provides for an installation to perform the process, wherein the installation comprises, fluidically connected in the following order: -a reactor 1 having two reactor sections (3. 5) and configured to transfer the heatgenerated in a reactor section to another reactor section;- an optional compression unit 27;- an optional purification unit 35;- an optional drying unit 43;- a hydrocarbon separation unit 51;- a CO removal unit 59;- a water-gas shift unit 73; and- an optional hydrogen separation unit 83;with preference, the reactor 1 is a shell-tube reactor comprising a first reactor section 3 beinga shell, and a second reactor section 5 comprising one or more inner tubes.In an embodiment, wherein the installation comprises both a compression unit 27 and apurification unit 35; the compression unit 27 and purification unit 35 are arranged in parallel,with the compression unit 27 being arranged to receive the effluent exiting the second reactorsection 5 and the purification unit 35 being arranged to receive the effluent exiting the firstreactor section 3.The process of the disclosure is also remarkable in that it recovers carbon monoxide (CO)which is a valuable product that can be used as a basic chemical in the production of many other compounds. As it will be seen in detail, the process according to the disclosure allowsfor an increase in the hydrogen production in a decarbonized process by performing a watergas-shift reaction on the CO-rich stream to produce a H2-rich stream.Reference is made to figures 1 and 2 which show embodiments of the process using differentarrangements of shell-tube reactors to perform the heat integration of the OCM and steamcracking reaction. Figure 1 shows a co-current shell-tube reactor 1 comprising two reactor sections (3; 5) defining a first reactor section 3 being a shell part and a second reactor section 5 being atubular part, the shell part surrounding the tubular part, wherein the tubular part comprisesone or more straight tubes. Figure 2 shows a mixed counter-current / co-current shell-tube reactor 1 comprising two reactor sections (3; 5) defining a first reactor section 3 being a shell part and a second reactorsection 5 being a tubular part, the shell part surrounding the tubular part, wherein the tubularpart comprises one or more U-shaped tubes. Whether the tubular part comprises U-shaped tubes or straight tubes, the tubular part may comprise a single tube or may comprise a set of two or more tubes.In both embodiments, the shell-tube reactor 1 comprises two inlets (7, 9) and two outlets (11,13) and the process comprises providing a first reactant mixture 15 comprising methane,oxygen, and optional carbon dioxide; and feeding the first reactant mixture 15 in the firstreactor section 3 of the shell-tube reactor 1 using a first inlet 7. The process comprisesperforming an Oxidative Coupling of Methane (OCM) reaction on the first reactant mixture in the first reactor section 3 of the reactor 1 to produce a first product stream 17 comprising carbon monoxide, ethylene and hydrogen that exits the reactor using a first outlet 11. Thefirst product stream 17 may also contain at least one selected from H2O, unreacted methane,and some minor amounts of ethane and CO2.The process also comprises providing a second reactant mixture 19 comprising hydrocarbonscontaining two or more carbon atoms in their molecular structure and feeding the secondreactant mixture 19 in the second reactor section 5 of the shell-tube reactor 1 using a secondinlet 9. The process comprises performing a steam cracking reaction on the second reactantmixture 19 to produce a second product stream 21 comprising ethylene and hydrogen that exits the reactor using a second outlet 13. It is understood that since the OCM reaction is exothermic, the heat generated by the OCM reaction on the first reactant mixture 15 in the first reactor section 3 is transferred to the secondreactant mixture 19 in the second reactor section 5 which to performs the endothermic steamcracking reaction. In the first arrangement illustrated in Figure 1, the first reactant mixture 15 flows co-currentlyrelative to the flow direction of the second reactant mixture 19 with the tubular part consistingof a straight tube.In the second arrangement illustrated in Figure 2, the first reactant mixture 15 flows counter-currently and co-currently relative to the flow direction of the second reactant mixture 19 in thefirst and second half of U-shaped tube, respectively. In both arrangements, the tubular part may comprise two or more tubes. In such a case,wherein one shell envelope contains multiple tubes, it is preferable that the open sectionalarea of the shell is between 80 and 120% of the sum of the sectional areas of all tubes. Themultiple tubes are typically fixed onto two plates or sheets, one on the inlet and one at theoutlet. A tube sheet is usually made from a round flat piece of plate sheet with holes drilled toaccept the tubes to support and isolate tubes in heat exchanger configuration. Hence the feedis homogeneously distributed over the multiple tubes while the shell side is fed laterally or viaa connected inlet pipe passing through the tube sheet.As it can be understood from the above, the reactor 1 is selected to have a cross-sectionalarea ratio between the second section and the first section (i.e. the tubular part and the shellpart) of about 0.8:1.2; preferably of about 0.9:1.1; more preferably of about 1.0:1.0. For example, the cross-sectional area of the second reactor section and first reactor section is the same.Also, the linear velocity of the flow between second section and the first section (i.e. the tubularpart and the shell part) of about 0.8:1.2; preferably, of about 0.9:1.1; more preferably, of about1.0:1.0. For example, the linear velocity of the flow within the tubular part and the shell part is the same. For example, the reactor is selected to have a tubular part with an internal diameter (Dt) of atleast 40 mm, a wall thickness between the tubular part and the shell part is about 2 mm, andthe shell part an internal diameter (Ds) of at least 80 mm; wherein the cross-sectional arearatio between the tubular part and the shell part of about 0.8:1.2. With the said dimensions thelength of the reactor is at least 3 m if the reactor comprises 2 inlets and at least 4 m if thereactor comprises 1 inlet.For example, the length and diameter of the tubular part of the reactor are selected to have aReynold number (Re) equal to or greater than 10,000 in each of the one or more tubes of thetubular part, where Re = uf DH ρf / μf, with uf being the velocity of the fluid, ρf density of the fluid, μf dynamic viscosity of the fluid, and DH hydraulic diameter of the tube (DH is equal to the innertube diameter) or shell (DH is equal to the inner shell diameter – outer tube diameter).In both arrangements, a person skilled in the art may find it advantageous to perform apreheating of the first reactant mixture 15 and / or of the second reactant mixture 19. Such a preheating step is advantageously performed in the case wherein the shell-tube reactor is aco-current shell-tube reactor. The preheating can be done by any means. In an embodiment,preheating can be done by a preheater 23 (such as a furnace heater, a heat exchangercoupled with the WGS reactor, fuel combustion heated, steam-fed heat exchanger, orelectrical heater). The OCM reaction Methane can be used to produce ethylene through the oxidative coupling of the methane (OCM) reaction. OCM reaction is known by the person skilled in the art and is described forexample in WO2018 / 146591A1, EP4082996, and WO2021 / 250002A2 which are incorporatedherein by reference.The OCM reaction in the present disclosure is a thermal OCM reaction; so the OCM reactionis performed in the absence of a catalyst. The fact that the reactor 1 operates without a catalyst increases the availability of the plant since there are no production losses due to the deactivation of the catalyst with time. Moreover, maintenance operations of a reactor without a catalyst are safer (no health risks caused by exposure of the operators to the catalyst dust) and faster. In addition, the fact that the reactor operates without a catalyst allows the production of a first product stream 17 containing not only ethylene and hydrogen but also carbon monoxide (as evidenced in Table 1 in the Example section).It is preferred that the reactor dimension and the flow of the gaseous feed stream be selectedto have a turbulent flow regime within the shell part of the reactor.In a thermal OCM reaction, it is possible to initiate the exothermic reaction at the start by over-preheating the feed with an elevated initial methane-to-oxygen ratio and / or by spark ignitionor plasma activation of the preheated feed. The spark ignition or plasma activation can be done at any location in the reactor.It is preferred that the process further comprises a start-up step of initiating the oxidativecoupling of methane reaction; wherein the start-up step comprises providing a first reactantmixture 15 comprising a gas mixture of methane (CH4) and oxygen (O2) in an initial molar ratioof at least 25.0 wherein the first reactant mixture 15 is preheated to an initial inlet temperatureof at least 650°C; with preference of at least 750°C. The light-off temperature (TLO) in the oxidative coupling of methane (OCM) reaction is the temperature at which the reaction begins to occur. It is the minimum temperature required toinitiate the reaction. In a non-catalytic OCM reaction, the light-off temperatures are preferablyabove 750 °C, more preferably above 850 °C.For example, in the start-up step, the first reactant mixture 15 is preheated to an initial inlettemperature of at least 650°C; preferably, at least 680°C; more preferably, at least 700°C;even more preferably, at least 750°C; and most preferably, at least 800°C.For example, in the start-up step, the first reactant mixture 15 is preheated to an initial inlettemperature ranging from 650 to 1200°C; preferably from 700 to 1100°C; more preferably from 750 to 1050°C and even more preferably, from 800 to 1000°C.For example, in the start-up step, the pressure in the reactor is at least 0.1 MPa; preferably atleast 0.2 MPa; more preferably at least 0.3 MPa, and even more preferably at least 0.4 MPa.In some embodiments, the pressure in the reactor is at least 0.5 MPa; preferably, at least 0.6MPa or at least 0.7 MPa.For example, in the start-up step, the pressure in the reactor is at most 2.0 MPa; preferably atmost 1.8 MPa; more preferably at most 1.5 MPa and even more preferably at most 1.2 MPaor at most 1.0 MPa. In some embodiments, the pressure in the reactor is at most 0.9 MPa, orat most 0.8 MPa, or at most 0.6 MPa.For example, in the start-up step, the pressure in the reactor ranges from 0.1 to 2.0 MPa;preferably from 0.2 to 1.8 MPa; more preferably from 0.3 to 1.5 MPa.Once the exothermic reaction is initiated, the methane-to-oxygen ratio can be slowly reducedwhile the preheating of the first reactant mixture 15 is reduced and the operating pressure isadjusted to the desired value. Once the OCM reaction is initiated the process comprises reducing both the inlet temperature and the molar methane to oxygen ratio. The inlet temperature is reduced from the initial inlet temperature to the operating inlet temperature. For example, the inlet temperature reduction is at least 100°C; preferably at least 150°C; more preferably at least 200°C. In some cases,the inlet temperature is reduced to at least 300°C or at least 400°C or at least 500°C or atleast 600°C or at least 700 °C. The reduction of the inlet operating temperature allows areduction in costs.For example, the process comprises preheating the first reactant mixture 15 to an operatinginlet temperature of at least 10 °C; preferably, at least 20°C; more preferably, at least 30°C; even more preferably, at least 40°C. In some embodiments, the operating inlet temperature isat least 200°C, or at least 250°C, or at least 300°C, or at least 350°C.For example, the process comprises preheating the first reactant mixture 15 to an operatinginlet temperature of at most 750°C; preferably, at most 650°C; more preferably, at most 600°C;even more preferably, at most 550°C; and most preferably, at most 520°C. In someembodiments, the operating inlet temperature is at most 300°C, or at most 250°C, or at most200°C.For example, the process comprises preheating the first reactant mixture 15 to an operatinginlet temperature ranging from 10 to 750°C or from 10 to 600°C; for example, from 30 to 550°C; for example, from 40 to 520°C. In some embodiments, wherein the selectivity to ethylene is to be favoured or the selectivity to CO2is to be kept as low as possible, the operating inlettemperature ranges from 200 to 550°C, or from 350 to 550°C. In some embodiments, whereinthe production rate of ethylene or the methane conversion rate is to be favoured, the operatinginlet temperature ranges from 10 to 300°C, or from 30 to 200°C.The molar ratio CH4:O2is reduced from the initial molar ratio CH4:O2to the operating molar ratio CH4:O2. The reduction of the molar ratio CH4:O2; i.e., the increase of the oxygenconcentration in the first reactant mixture 15, allows an increase of the selectivity to C2+products. It is understood that, in the context of the disclosure, the terms “C2+ products” referto hydrocarbons that contain two or more carbon atoms in their molecular structure.For example, the first reactant mixture 15 shows an operating molar ratio CH4:O2 of at least2.5; preferably at least 3.0 or at least 3.5, more preferably at least 5.0, even more preferablyat least 6.5 and most preferably at least 7.0 or at least 7.5.For example, the first reactant mixture 15 shows an operating molar ratio CH4:O2 of at most20.0; preferably at most 18.0, more preferably at most 15.0, and even more preferably at most14.0. In some embodiments, the first reactant mixture 15 shows an operating molar ratioCH4:O2 of at most 10.0, or at most 9.5, or at most 9.0, or at most 8.5.For example, the first reactant mixture 15 shows an operating molar ratio CH4:O2 ranging from2.5 to 20.0; preferably, from 3.0 to 18.0; more preferably from 3.5 to 15.0, and even morepreferably from 5.0 to 10.0. In some embodiments the operating molar ratio CH4:O2 rangesfrom 6.5 to 9.5 or from 7.0 to 9.0.For example, the first reactant mixture 15 further comprises one or more hydrocarbonsdifferent from methane; with preference, the first reactant mixture 15 further comprises ethane. It is preferred that the ratio of the volume (V) of the reactor divided by the flow rate (FCH4,0) ofmethane in the feed stream (i.e., V / FCH4,0) is at most 1.0 m3s mol-1; preferably, at most 0.8 m3smol-1; more preferably, at most 0.5 m3s mol-1; even more preferably at most 0.2 m3s mol-1 andmost preferably, at most 0.1 m3s mol-1. For example, the ratio of the volume (V) of the reactor divided by the flow rate (FCH4,0) ofmethane in the feed stream (i.e., V / FCH4,0) is at least 0.002 m3s mol-1; preferably at least 0.005m3s mol-1 or at least 0.008 m3s mol-1. For example, the process comprises an operatingpressure in the reactor of at least 0.1 MPa; preferably at least 0.2 MPa; more preferably atleast 0.3 MPa, and even more preferably at least 0.4 MPa. In some embodiments, theoperating pressure in the reactor of at least 0.5 MPa, or at least 0.6 MPa, or at least 0.7 MPa.For example, the process comprises an operating pressure in the first reactor section of thereactor is at most 5.0 MPa; preferably at most 4.0 MPa or at most 3.5 MPa; more preferablyat most 3.0 MPa and even more preferably at most 2.5 MPa; most preferably of at most 2.0MPa or at most 1.0 MPa. In some embodiments, the operating pressure in the reactor is atmost 0.9 MPa, or at most 0.8 MPa, or at most 0.6 MPa.For example, the process comprises an operating pressure in the first reactor section rangingfrom 0.1 to 4.5 MPa; preferably, from 0.1 to 4.0 MPa; preferably, from 0.1 to 3.5 MPa;preferably, from 0.1 to 3.0 MPa; preferably, from 0.1 to 2.5 MPa; preferably, from 0.1 to 2.0MPa, preferably, from 0.2 to 1.8 MPa, preferably, from 0.2 to 1.0 MPa, preferably, from 0.3 to 1.5 MPa, preferably, from 0.3 to 0.9 MPa, preferably, from 0.4 to 0.8 MPa.In an embodiment, the first product stream 17 comprises methane, and methane isseparated from the first product stream 17 and recycled in the first reactant mixture 15 to bereintroduced in the first reactor section 3. For example, methane is separated from the firstproduct stream 17 by distillation and the installation further comprises one or more distillation columns.The steam-cracking reactionThe steam cracking reaction is performed in the absence of a catalyst.The steam cracking reaction is preferably performed on a second reactant mixture 19 comprising one or more selected from ethane, propane, butane, liquefied petroleum gas,naphtha, and diesel; with preference, the second reactant mixture 19 comprises ethaneand / or propane. Liquefied petroleum gas (LPG) comprises mainly propane and butanes. Petroleum naphthaor naphtha is defined as the hydrocarbon fraction of petroleum having a boiling point from15°C up to 200°C. It is a complex mixture of linear and branched paraffins (single and multi-branched), cyclic paraffins, and aromatics having carbon numbers ranging from 5 to about 11carbon atoms. Light naphtha has a boiling range from 15 to 90°C and comprises C5 to C6hydrocarbons, while heavy naphtha has a boiling range from 90 to 200°C and comprises C7to about C11 hydrocarbons. Diesel has a boiling range from about 200 to 350°C and compriseC10 to C22 hydrocarbons, including essentially linear and branched paraffins, cyclic paraffins,and aromatics (including mono-, naphtho- and poly-aromatic). Heavier diesel (likeatmospheric gasoil, vacuum gasoil, atmospheric residua, and vacuum residua), having boilingranges above 300°C and C20+ hydrocarbons including essentially linear and branchedparaffins, cyclic paraffins, and aromatics (including mono-, naphtho- and poly-aromatic) areavailable from atmospheric or vacuum distillations units. For example, the process comprises preheating the second reactant mixture 19 to anoperating inlet temperature of at least 100 °C; preferably, at least 200°C; more preferably, atleast 300°C; even more preferably, at least 400°C. In some embodiments, the operating inlettemperature is at least 500°C, preferably at least 550°C, preferably at least 600°C, preferablyat least 650°C.For example, the process comprises preheating the second reactant mixture 19 to anoperating inlet temperature of at most 700°C, preferably at most 600°C, preferably at most550°C; preferably, at most 520°C. In some embodiments, the operating inlet temperature is at most 300°C, preferably at most 250°C; preferably at most 200°C.For example, the process comprises preheating the second reactant mixture 19 to anoperating inlet temperature ranging from 10 to 700°C, preferably from 10 to 600°C; forexample, from 30 to 550°C; for example, from 40 to 520°C.The steam cracking reaction is preferably conducted at a temperature ranging from 500°C to1200°C, preferably, ranging from 700°C to 1000°C; more preferably, ranging from 750°C to900°C. In a preferred embodiment, the steam cracking reaction is performed in the absenceof catalyst and is conducted at a temperature ranging from 750°C to 900°C.Control that sufficient heating is provided to the steam cracking reaction is made by controllingthe flow rate of methane and oxygen in the OCM shell so that the OCM reaction providesenough calories for steam cracking. For example, the outlet temperature of the reactor mayrange from 800 to 1,200°C, preferably from 820 to 1,100°C, more preferably from 825 to1,000°C, more preferably from 830 to 950°C, more preferably from 840°C to 900°C. For example, the steam cracking reaction is performed at a pressure ranging between 0.1 MPa and 1.0 MPa, preferably between 0.1 MPa and 0.5 MPa.The residence time of the second reactant mixture 19 in the reactor where the temperatureis between 500 and 1,200°C, which may advantageously range from 0.005 to 1.00 seconds,preferably from 0.01 to 0.60 seconds, more preferably from 0.10 to 0.30 seconds, even morefrom 0.10 to 0.25 seconds. Such low residence time is advantageous to avoid secondary reactions and thus prevent the formation and deposition of coke.The second product stream 21 comprises the cracking products and unreacted products. Inparticular, the cracking products obtained may include one or more from ethylene, propyleneand benzene; and optionally one or more from hydrogen, toluene, xylenes, and 1,3-butadiene.In an embodiment, the second product stream 21 comprises methane, and methane isseparated from the second product stream 21 and recycled in the first reactant mixture 15.The optional compression step performed on the second product streamIn one or more embodiments, the pressure of the first product stream 17 is higher than thepressure of the second product stream 21, and the process comprises a compression stepwherein the second product stream 21 is compressed to at least the pressure of the first product stream 17 and subsequently mixed with the first product stream 17 to form a mixed stream 25 before the step of separating at least a part of the carbon monoxide contained in,or originating from, the first product stream 17 to produce a CO-rich stream 71.However, in a preferred embodiment, the process is conducted on compressed gas, so the second product stream 21 is compressed and a compressed gas stream 33 is obtained. For example, the compressed gas stream 33 has a pressure ranging from 0.5 to 5.0 MPa; preferably from 1.0 to 4.0 MPa; more preferably from 1.2 to 3.8 MPa, and even more preferably from 1.5 to 3.6 MPa. The compression step may comprise a quenching sub-step which is preferably performed before the compression.When a substep of compression is conducted the installation comprises a gas compressorunit 27 placed upstream of the CO-removal unit; wherein the gas compressor unit comprisesone or more gas compressors 29 and optional quenching means 31.When a substep of compression is conducted, the downstream streams are understood to becompressed stream unless indicated differently.The optional purification stepAt least one of the first product stream 17, the second product stream 21, or the compressedgas stream 33, can be subjected to a step of purification. With preference, as illustrated onFigure 6, at least the first product stream17 is subjected to a step of purification.The fact that the reactor 1 operates without a catalyst to perform the OCM reaction allows theproduction of a first product stream 17 containing carbon monoxide, ethylene, and hydrogenbut also carbon dioxide (as evidenced in Table 1 in the Example section).The purification step is conducted to remove acid gases such as CO2 and optional H2S.Removal of CO2is beneficial to improve the CO conversion in the later step of performing a water gas shift reaction.With preference, the purification step is or includes an amine wash and / or a caustic wash.Acid gas (such as carbon dioxide) removal can be made by any means; for example, the purification step can be done in a single step or two or more sub-steps. In a preferred embodiment, the purification step is a two-sub-step procedure. In that respect, the installationto perform the gas product separation process comprises a purification unit 35.In any case, it is preferred that purification is a step or comprises a sub-step of purification byamine solution (i.e. amine wash). The same is preferably performed in a purification unitcomprising an amine absorption column 37. CO2-removal by amine solution is well-known tothe person skilled in the art and is described for example in Dubois et al. in “Carbon dioxideabsorption into aqueous amine-based solvents: modeling and absorption tests”; EnergyProcedia 4, 1353–1360 (2011), which is incorporated herein by reference.In an embodiment, the amine solution comprises one or more amines selected frommonoethanolamine (MEA), diethanolamine (DEA), diglycolamine (DGA),methyldiethanolamine (MDEA), triethanolamine (TEA), piperazine (PZ) and (piperazinyl-1)-2- ethylamine (PZEA); preferably selected from monoethanolamine (MEA), diethanolamine (DEA), piperazine (PZ) and (piperazinyl-1)-2-ethylamine (PZEA); more preferably, the aminesolution comprises one or more amines selected from monoethanolamine (MEA),diethanolamine (DEA), and piperazine (PZ); even more preferably, the amine solutioncomprises monoethanolamine (MEA) and / or diethanolamine (DEA); and most preferably, theamine solution comprises monoethanolamine. Indeed, monoethanolamine was found to showhigh reactivity with CO2.For example, the amine solution comprises piperazine (PZ) at a concentration ranging from 5to 35 wt.% based on the total weight of the solution; preferably from 10 to 20 wt.%.In a preferred embodiment, the amine solution comprises monoethanolamine (MEA) at aconcentration ranging from 10 to 40 wt.% based on the total weight of the amine solution; preferably from 10 to 20 wt.%.Advantageously, the purification unit comprises an amine absorption column 37, where CO2is removed, and a regeneration (stripper) column (not shown), where the chemically bondedCO2is released in a separate gas stream on top of the column. The regenerated aminesolution is sent back to the absorption column through a recycle loop.Having in mind that the CO2absorption with amines involves chemical reactions, higher pressure favours kinetics of the CO2absorption, and results consequently in a lower amountof the amine that circulates in the purification unit 35. However, higher pressure translates intoa higher cost of the gas compression section upstream of the CO2absorber and a higher costof the construction material. Thus, the CO2 removal by amine is preferably conducted at apressure ranging from 0.5 to 5.0 MPa; preferably from 1.0 to 4.0 MPa; more preferably from1.2 to 3.8 MPa, and even more preferably from 1.5 to 3.6 MPa. For example, an operatingpressure of the CO2 absorption column ranging from 1.2 to 2.0 MPa results in acceptable flow rates of the amine solution with acceptable gas compression cost upstream the column.Since amines in water solution are weak bases, there may be some remaining portion of CO2present in the gas phase after CO2 removal with the amine solution, typically in the range of50-500 ppm CO2. Therefore, the CO2 removal step by amine is preferably followed by a CO2removal step with an alkaline solution (also named caustic washing), such as with a solutionof KOH and / or NaOH (also named caustic soda) in a caustic tower 39. The caustic tower 39consists of two packed beds. The top bed operates with circulating wash water to remove any residual caustic entrained with the vapors in the gas stream from being carried over to downstream equipment while the bottom bed operates with circulating caustic.Advantageously, the base used is strong.The purification step aims to obtain a purified stream 41 (i.e. a CO2-lean gas stream) thatcontains less than 5 molar ppm of CO2 based on the total molar content of the CO2-lean gasstream, which is desirable to have downstream a cryogenic separation column.Thus, in an embodiment, the purification step (CO2-removal) is a step or comprises a sub-stepof caustic washing; with preference, the alkaline solution is selected from a sodium hydroxide(NaOH) solution, a calcium hydroxide solution (Ca(OH)2), a potassium hydroxide (KOH)solution, and any mixture thereof. For example, the alkaline solution comprises from 10 to 35 wt.% of a base based on the totalweight of the alkaline solution, wherein the base is selected from sodium hydroxide, calciumhydroxide, potassium hydroxide, and any mixture thereof; preferably the base is sodiumhydroxide and / or potassium hydroxide.In a preferred embodiment, the step of purification comprises a sub-step of acid gas removal(such as CO2 removal) by an amine solution followed by a sub-step of acid gas removal (suchas CO2 removal) by an alkaline solution. In other words, the step of purification comprises asub-step of amine washing followed by a sub-step of caustic washing.For example, the purification step is performed in two sub-steps:1) chemical absorption by an amine solution wherein the amine solution comprises from10 to 40 wt.% of monoethanolamine based on the total weight of the amine solution, and 2) removal of the remaining acid gases (such as CO2) with chemical absorption by an alkaline wherein the alkaline solution comprises from 10 to 35 wt.% of sodiumhydroxide based on the total weight of the alkaline solution.At the end of the purification step a purified stream 41 (i.e. a sweet gas stream) is obtainedThe optional step of hydrocarbon recoveryIn a preferred embodiment, a step of hydrocarbon recovery is performed before the step ofseparation of CO. In a preferred embodiment, the hydrocarbon recovery is performed by cryogenic distillation. Itsuch a case, the step of hydrocarbon recovery includes sub-steps of drying and cooling thestream. In a preferred embodiment, the sub-steps of drying and cooling are performed on thefirst product stream 17 (or the purified stream 41 if applicable), or on the second productstream 21 (or the compressed stream 33 if applicable); or on the mixed stream 25.The purpose of driers in the drying step (i.e. the H2O removal step) is to remove the water byadsorption on the adsorbent bed to ensure that the stream has less than 1 weight ppm waterbefore the cryogenic distillation. Water can freeze in the cold sections affecting columnperformance and causing maintenance issues.This drying step may be accomplished by using one of several methods. Examples aretriethylene glycol contacting, membrane permeation, or adsorption with a regenerableadsorbent such as a small pore molecular sieve. A suitably designed drying unit 43 will drythe stream to a few parts per million water.Regenerant is heated in the regenerant heater to about 230 °C using temperature control andcontacts the adsorbent in up flow direction. A shell & tube steam heater or an electric heater or a combination of both may serve as the heater. If a steam heater is used, the temperature is controlled by regulating the condensate flow. In the case of an electric heater, the temperature is controlled by regulating the power to the electric heater bundle elements. The regenerant effluent from the drier is cooled in the regenerant cooler and flows to the regenerant knockout drum. Any liquids collected in the regenerant knockout drum are removed using the level control. Regenerant effluent from the regenerant knockout drum is then routed to the downstream system. The regeneration system pressure will be floating on the back pressure of the downstream system depending on where the regenerant effluent is routed. Regenerant effluent may be routed to the fuel gas system or hydrogen purification system (PSA). After the heating step, the drier is cooled and readied to be placed online. The drier that has been regenerated is placed online while the other drier is taken offline for regeneration.In a preferred embodiment, the sub step of drying consists of or comprises using a vapor-liquid separator 45.In an embodiment, the substep of drying consists of or comprises using small pores molecularsieves. Molecular sieves are crystalline metal aluminosilicates having a three-dimensionalinterconnecting network of silica and alumina tetrahedra. The critical pore diameter value is3.2 Å. Small pores molecular sieves with a pore diameter of 3 Å are commercially availablefrom UOP or BASF.In a preferred embodiment, the sub-step of drying comprises a first sub-step of vapor-liquidseparation or adsorption, followed by a second sub-step of drying using small pores molecularsieves. For example, the drying unit 43 comprises a vapor-liquid separator 45 followed by a molecularsieves dryer system 47. In another example, the drying unit comprises an adsorption systemfollowed by a molecular sieves dryer system; with preference, the adsorption system comprises a plurality of hollow fibers.A dried gas stream 49 (i.e. a dehydrated gas stream) exits the drying unit 43 and is directedto a hydrocarbon separation unit 51 for hydrocarbon separation.The hydrocarbon separation unit 51 is preferably a cryogenic unit that comprises one or morecolumns 53, ethylene and other C2+ hydrocarbons are separated from the rest of the gases.At the bottom of the column, C2+ products (including ethylene) 55 are collected. At the top a gaseous phase 57 comprising CO, methane and hydrogen is collected.The dried gas stream 49 is introduced into the cryogenic column 53 where it is progressivelycooled to very low temperatures using refrigeration systems or turboexpander technology. Asthe temperature drops, the heavier hydrocarbons (C₂+) condense into a liquid phase whilemethane and other light gases such as CO remain in the gaseous phase and are collected atthe top of the column.The step of separating at least a part of the carbon monoxide contained in, or originating from,the first product stream 17 to produce a CO-rich streamAccording to the disclosure, the process comprises an optional step of separation of the COto obtain a CO-rich stream 71. The gaseous phase 57 exiting the cryogenic separation isconverted to a CO removal unit 59 wherein the CO is separated from methane and hydrogen.The step of removal of carbon monoxide can be done by any method known to the personskilled in the art. For example, it can be performed by cryogenic distillation, by pressure swingadsorption (PSA), or by the COSORB process.Cryogenic distillation utilizes the difference in boiling points of hydrogen, methane and CO. In a preferred embodiment, hydrogen is first separated from CO and CH4and then sequentialcooling and fractionation separate CO and CH4. As CH4 is less volatile it condenses at highertemperature compared to CO which remains in the gas phase with hydrogen.Alternatively, as illustrated in Figure 6, CH4 is separated first, and CO and H2 in a subsequentstep. In such a case, the CO removal unit 59 may comprise a cryogenic separation column 61to separate a CH4-rich stream 63 from a gaseous stream 65 which is further submitted to ahydrogen separation substep. Hydrogen separation is preferably performed in an H2-removal sub-unit 67 comprising one or more H2 separation membranes. Suitable H2 separation membranes are known to the person skilled in the art. For example, suitable H2 separation membranes comprise palladium or palladium alloys and are described in US8747766B2 which is incorporated herein by reference. For example, the H2 separation membrane may comprise an alloy comprising a Group 5 element selected from V, Nb, Ta, or a combination thereof and Ir. Such an H2separation membrane is disclosed in EP2596851B1 which is incorporated herein byreference. After hydrogen separation, a first H2 rich stream 69 is recovered as well as a CO-rich stream 71. When separation is performed by pressure swing adsorption, the installation comprises a PSA device. For example, the PSA device comprises one or more adsorbers being one or more zeolites selected from zeolite 4A, zeolite 5A or zeolite 13X.COSORB process exploits the ability of cuprous ions (Cu+) to form a reversible complex withCO: Cu++CO ↔ Cu(CO)+The COSORB process is described in Hogendoorn et al. in “The absorption of carbonmonoxide in COSORB solutions: absorption rate and capacity” The Chemical EngineeringJournal 59 (1995) 243-252 (doi: 10.1016 / 0923-0467(94)02959-8) which is incorporated byreference.As a result of the step of CO recovery, a CO-rich stream 71 is obtained and sent to a water-gas shift unit 73. A methane-rich stream 63 is also recovered and can be at least partlyrecycled in the first reactant mixture 15 as a feedstream for the OCM reaction. A first hydrogenstream 69 is also recovered.The step of performing a water gas shift reaction on the CO-rich streamThe water gas shift (WGS) reaction allows carbon monoxide to be converted into carbondioxide and hydrogen. The shift reaction is performed in water gas shift reactors and is well known to a person skilled in the art. Suitable processes to produce hydrogen from afeedstream containing CO are described in US20040081614 and WO2022 / 017829, whichare incorporated by reference. The WGS reaction is carried out at a temperature above 200°C, preferably ranging from 200 to 500°C, more preferably from 300 to 500°C, most preferably from 350 to 450°C.The WGS reaction is preferably carried out at a pressure ranging from 1.0 to 4.0 MPa.A water stream 79 is provided to the one or more reactors preferably in a steam form. The steam-to-carbon ratio preferably ranges from 1.1:1 to 5:0, preferably from 2:1 to 3:1 as an excess of H2O shifts the equilibrium of the reaction towards H2 production. The WGS reaction is exothermic, the conversion of CO is increased with lower temperature.With preference, the process comprises transferring a part of the heat generated by the watergas-shift reaction to pre-heat the first reactant mixture 15 and / or the second reactant mixture19 before entering the reactor 1 having two reactor sections to perform the coupling OCM-SC. The WGS reaction can be performed in a single reactor such as the second reactor 75illustrated in Figure 6) or in at least two reactors arranged in series (not illustrated).For example, the WGS reaction can be performed in at least two reactors in series whereinthe reaction is conducted at different temperatures. The reactor with the higher temperatureis named the high-temperature shift reactor (i.e. HTS reactor) to perform an HTS reaction,and the reactor with the lower temperature is named the low-temperature shift reactor (i.e.LTS reactor) to perform an LTS reaction. The HTS reactor being placed upstream the LTSreactor, so the HTS reaction is performed before the LTS reaction. The use of such a configuration with two reactors in series is found to increase the overall conversion of CO. In an embodiment wherein the purification step (i.e. the CO2removal) has not been performed upstream of the WGS reaction, the WGS reaction consists of the HTS reaction if the WGS reaction is conducted in a single reactor or comprises a HTS reaction when the WGS reaction is conducted in two reactors in series (with the HTS reaction be performed in the first reactor, and the LTS reaction be performed in the second reactor). In a preferred embodiment wherein the purification step (i.e. the CO2 removal) has been performed upstream of the WGS reaction, the WGS reaction can be performed in a single reactor at low temperature (corresponding to the LTS reactor). In other words, in the absenceof CO2 in the feedstream (i.e. in the CO-rich stream), the LTS reaction only is performed.For example, the HTS reaction is performed at a temperature ranging more than 300°C to 500°C; preferably from 310 to 450°C. For example, the LTS reaction is performed at a temperature ranging from 200 to 300°C;preferably from 200 to 280°C and more preferably from 200 to 250°C.It is preferred that the WGS reaction (i.e. the HTS and / or LTS reaction) be performed in thepresence of both a catalyst and an adsorbent for the adsorption of at least CO2. With preference, the water-gas shift reaction and said adsorption of CO2 on an adsorbent areperformed in the same reactor so both the catalyst and the adsorbent are loaded in the WGSreactors (i.e. in the HTS and LTS reactors). The use of an absorbent in the WGS reactorallows for enhanced CO conversion beyond equilibrium limits.In an embodiment, the adsorbent and the water-gas shift catalyst are different materialsloaded in the same reactor, wherein preferably the weight ratio of the adsorbent to catalyst is in the range of about 2-50, such as about 5-20, especially about 10-20, or in the range of about 20-100, such as 20-50, especially about 25-50. A suitable catalyst for the HTS reaction is iron oxide promoted with chromium oxide (Fe3O4 / Cr2O3). Suitable catalysts for the LTS reaction are selected from copper oxide promoted withzinc oxide (CuO / ZnO) and copper oxide promoted with zinc oxide and alumina (CuO / ZnO-Al2O3).Adsorbents for CO₂ removal in the water-gas shift reaction include alkali-based sorbents suchas calcium oxide (CaO), magnesium oxide (MgO), and lithium silicate (Li4SiO4); zeolites such as 13X and 5A; activated carbon; amine-functionalized materials like amine-grafted silica and alumina; metal-organic frameworks (MOFs) such as MIL-101 and HKUST-1; hydrotalcites; lithium zirconate (LI2ZrO3); sodium-based sorbents like sodium oxide (Na2O) and sodium carbonate (Na2CO3); silica-based sorbents such as SBA-15 and MCM-41; metal oxides like zinc oxide (ZnO) and copper oxide (CuO); and dual-functional sorbents combining catalytic and adsorptive properties. For example, the adsorbent used in the WGS reaction comprises at least one selected from calcium oxide (CaO) and magnesium oxide (MgO).In an embodiment, the catalyst and the adsorbent are part of the same material loaded inthe same reactor.In such a case the adsorbent for shifting the WGS reaction towards hydrogen is preferably amaterial that is capable of adsorbing both H2O and CO2and preferably also H2S and othergaseous acidic components if present. A suitable adsorbent comprises, for instance, alkali-promoted alumina. The adsorbent to be used in the process of the disclosure is preferably an inorganic oxide, which comprises a trivalent metal oxide, in particular alumina, alumina oxide hydroxide or aluminum hydroxide. Instead of or in addition to aluminum, other metals capable of adopting a trivalent state may be present, such as Fe, Mn, Cr, Ti and Zr. The adsorbent furthermore comprises one or more alkali metal oxides, hydroxides and / or carbonates. Any alkali metal can be used, including Li, Na, K, Rb and Cs. Preferred alkalimetals are Na and K. For example, the molar ratio of alkali metal to the metal of said one ormore trivalent metal oxides is ranging between 0.10 and 1.00, preferably between 0.20 and 0.90, more preferably between 0.25 and 0.75.The adsorbent may advantageously further comprises one or more divalent metal oxides,hydroxides and / or carbonates. The divalent metals can be an alkaline earth metal (Mg, Ca, Sr, Ba) or Co, Ni, Cu, Zn, Cd, Pb. Preferred divalent metals are Mg, Ca, Sr, Ba, Zn, Ni and Cu; more preferably selected from Mg, Ca and Zn. More preferably, the adsorbent comprises calcium oxide and / or magnesium oxide and / or zinc oxide. For example, the molar ratio of divalent metals to the metal of said one or more trivalent metal oxides is ranging between 0.25 and 1.50, more preferably between 0.30 and 1.25, even more preferably between 0.40 and 1.00. In particular, the adsorbent has a molar (= atomic) ratio of divalent metals (especially one or more of Mg, Ca, Zn) to Al of between 0.25 and 1.50, preferably between 0.4 and 1.0 and an atomic ratio of alkali metal (especially Na and / or K) to Al of between 0.1 and 1.0, preferably between 0.25 and 0.75. Where higher WGS catalyst activity is required, the presence of zinc was found to be advantageous, for example using a Zn: Al molar ratio of 2:98 to 80:20, especially 5:95 to 60:40, within the above ratios of total divalent metals to aluminum. Preferably, the adsorbent comprises magnesium oxide (magnesia) and has a molar (atomic) Mg to Al+Mg ratio of between 0.05 and 0.95, more preferably between 0.1 and 0.8, most preferably between 0.2 and 0.6. Where reference is made to alumina, magnesia and the like, these include the oxides, but also hydroxides and other equivalents of the oxides of aluminum, magnesium, respectively. Magnesium is particularly preferred for feed gas mixtures containing significant amounts of sulphur-containing contaminants such as H2S, as further detailed below, since the magnesium-based adsorbents were found to be insensitive to the sulphur compounds. Aluminas also containing alkali metals, possibly in addition to other metals and counterions, are referred to herein as "alkali-promoted aluminas". Aluminas, also containing magnesium and / or other divalent metals, and also containing alkali metals, possibly with other metals and counterions, are referred to herein as "alkali-promoted hydrotalcites". The aluminas may be used in a manner which may comprise admixing metals oxides and further additives with the alumina or hydrotalcite or other base material in a dry state or in a solution or in a slurry, and optionally drying and calcining the resulting mixture. The alumina may be any form of alumina that can be rehydrated, in particular, which has a level of hydroxyl groups. Examples includegamma-alumina, boehmite, gibbsite, and bayerite.More generally, inorganic oxides which can be used as an adsorbent can be represented by the following chemical formula: [Mg((1-x )β) MII((1-x)(1- β)) Al(αx) MIII((1-α)x)(OH)y] [Zn-]((x-y+2) / n) p H2OqMI(Am-)1 / m, wherein: MIis one or more metals selected from Li, Na, K, Rb and Cs; preferably selected from Na and K; MIIis one or more metals selected from Ca, Sr, Ba, Co, Ni, Cu, Zn, Cd and Pb, preferably Ni, Cu, Zn; more preferably Zn; MIIIis one or more metals selected from Fe, Mn, Cr, Ti and Zr; Zn-is one or more anions selected from halide, nitrate or acetate (n=1), or oxide, sulphate, oxalate or carbonate (n=2); Am-is one or more anions selected from hydroxide (m=1) and the anions as defined for Z above, with m corresponding to n; m and n = 1 or 2 according to A and Z, respectively; x = 0.05-1, preferably 0.05-0.95, more preferably 0.20-0.90; α = 0-1, preferably 0.5-1 ; β = 0-1 preferably 0.25-1, more preferably 0.5-1 ; p = 0-15; q = 0.1-1; y = 0-4. For example, the adsorbent used is or comprises an inorganic oxide represented by the chemical formula [MII(1-x) Al(αx) MIII((1-α)x)(OH)2][Cn-](x / n) y (H2O)z MI(m) Am-wherein MIis one or more alkali metals selected from Li, Na, K, Rb and Cs; preferably selected from Na and K; MIIis one or more divalent metals selected from Mg, Mn, Cu, Co, Fe, Cd and Cr; MIIIis optionally one or more trivalent metals selected from Fe, Cr and Mn; Cn-, wherein n = 1, is one or more anions selected from the group comprising halide, nitrate or acetate; wherein n = 2, is one or more anions selected from the group comprising sulphate or carbonate; Am-, wherein m =1, is one or more anions selected from hydroxide, halide, nitrate or acetate, or wherein m = 2, is one or more anions selected from sulphate, oxalate or carbonate; and wherein x, α, y and z are coefficients, wherein 0.01 ≤ x ≤ 0.99; preferably 0.3 ≤ x ≤ 0.7; 0.5 ≤ α ≤ 1;0 ≤ y ≤ 4; and 0.001 ≤ z ≤ 7; preferably 0.033 ≤ z ≤ 0.17. Where the adsorbent contains aluminum and magnesium and / or another divalent metal M n such as Ca or Zn, the proportion of aluminum, expressed by x in the above formula, is preferably relatively high, i.e. x = 0.45-0.90, especially when high pressures (> 1.5 MPa) are applied. At low pressures, e.g. below 1.5 MPa, in particular below 1.0 MPa, higher MII: Al ratios may be useful, e.g. x = 0.20-0.70. At intermediate pressures, x may be e.g.0.3-0.8. Specific examples of hydrotalcites of the above formula are referred to herein as KMG30 having a MgO: Al2O3weight ratio of 30:70 and having the formula [Mg0.35Al0.65(OH)2][CO32-]0.325*0.5H2O 0.32K(CO32-)0.5with a molar ratio K: Mg: Al of about 1.0: 1.1: 2.0; and as KMG50 having a MgO: Al2O3weight ratio of 50:50 with a molar ratio K: Mg: Al of about 1.0: 1.7: 1.4, and having the formula [Mg0.55Al0.45(OH)2][CO32-]0.2250.5H2O-0.32K(CO32-)0.5. The anions in the complex metal oxides preferably comprise hydroxide and / or carbonate anions to ensure sufficient alkalinity for effective adsorption of acidic gas species. In particular, at least 50% of the anions (expressed in monovalent equivalents) comprise hydroxide and / or carbonate. Suitable inorganic oxides can have a layered structure, wherein part of the anions is arranged in layers interposed between layers containing the cations. Examples of suitable layered oxides include the hydrotalcites having proportional formulas such as Mg6Al2(OH)16 (CO3) 4H2O or similar combinations with different Mg: Al ratios. Other suitable oxides include analogues wherein magnesium is absent (e.g. scarbroite) or is replaced by calcium (e.g. alumohydrocalcites), strontium (e.g. montroyalite) or barium (e.g. dreserrites), as well as Mg / Fe, Mg / Cr, Mg / Mn, Ni / Al etc. analogues (pyroaurite, stichtite, desautelsite, takovite). The adsorbent preferably contains an alkali metal compound. Such alkali-containing materials are referred to also as 'alkali-promoted'. Thus, the base material of the adsorbent can be alkali- promoted alumina. The alkali promoters may be in the form of oxides, hydroxides or, preferably carbonates. Especially, the alkali content is > 5 wt.% calculated as alkali metal, preferably 5-30 wt.%, relative to the final mixed oxide composition. The adsorbent may have been thermally treated, i.e. it may have been heated at a temperature above about 200°C, even more especially above about 400°C. For instance, assuming a hydrotalcite, when heating this hydrotalcite in the reactor before the WGS reaction or during the WGS reaction, the hydrotalcite modifies to a promoted alumina, such as K2CO3 and MgO promoted alumina, since, at elevated temperatures, the hydrotalcites may at least partially rearrange in mixed oxides while losing hydrotalcite crystalline structure and layered double hydroxide structure. This is well known in the art and is for instance described in US 5,358,701,US 6,322,612, and WO 2005 / 102916.At the end of the WGS reaction, an effluent stream 81 is obtained. In a preferred embodimenta hydrogen separation step is performed to recover an H2-rich stream 87 is recovered. CO2can be recovered to form a CO2-rich stream 85 and optionally recycled into the OCM reactor.With preference, the hydrogen separation step is conducted by pressure swing adsorption(PSA). Thus, to perform the hydrogen separation step, the installation may further comprise a hydrogen separation unit 83 downstream the water-gas shift unit 73; with preference the hydrogen separation unit 83 comprises one or more H2separation membranes and / or a PSA device.It can be understood that the process of the disclosure has a heat integration for a greenerproduction of ethylene together with improved hydrogen production. Heat integration is foundwith the coupling of the OCM and SC reaction in a single reactor and optionally in preheatingthe feedstream for the OCM and / or SC reactions by the heat generated by the WGS reaction.Hydrogen production is improved by the recovery of the CO from the OCM reaction that issubjected to the WGS reaction, so the overall integrated process produces hydrogen in bothsteps of SC and WGS by comparison to an integrated process devoid of the WGS step.The hydrogen produced in the steam cracking and the hydrogen produced in the OCM reaction combined with the hydrogen produced in the WGS reaction can optionally be used to heat the other stream cracking furnaces by combustion of the hydrogen, either alone or mixed with other fuels. ExamplesExample 1In this example, the oxidative coupling of methane (OCM) is combined with ethane steamcracking (SC) in a shell and tube reactor, where the tube has a U-shape. The heat generated by the oxidative coupling of methane (shell) is used in the ethane steam cracking that takes place in the inner U-tube (Figure 3). Both streams are preheated before entering the reactor. The temperature profile along the reactor can be observed in Figure 4. The conversions of methane in the shell and ethane in the tube are 16% and 57%, respectively. The compositions of inlet and outlet streams are described in Table 1. Table 1. Composition of inlet and outlet streams. Inlet shell Outlet shell inlet tube Outlet tubeComponent kg / h kg / h Component kg / h kg / hO2 8 032.83 51.81 H2O 393.36 393.78H2 439.17 H2 71.14H2O 5833.19 CH4 56.30CH4 36247.17 30476.74 C2H6 1966.80 835.40CO 4520.53 C2H4 910.56CO2 251.49 C2H2 7.57CH3OH 0.31 C3H8 2.43CH2O 1.56 C3H6 21.02C2H6 180.18 C3H4 0.42C2H4 1693.87 C3H4 0.88C2H2 136.62 C4H10 11.49CH3CHO 1.35 C4H8 4.37CH2CO 19.78 C4H6 28.43CHOCHO 29.43 C4H4 1.36CHCHO 28.25 C5+ 15.01C3H8 4.69 C3H6 165.32 CH3CHCO 3.47C3H4 77.88 C2H3CHO 0.54C4H8 15.98 C4H655.49 C4H4 40.21 C5+252.12 Example 2 This example is related to the process of the disclosure up to the CO-recovery step, i.e. before the WGS step.Reference is made to Figure 5. The n# reference in Figure 5 corresponds to the columns ofthe below table 2.

Claims

CLAIMS 1. A process to produce ethylene and hydrogen characterized in that it comprises:- providing a first reactant mixture (15) comprising methane, oxygen, andoptional carbon dioxide; -performing an Oxidative Coupling of Methane (OCM) reaction on the firstreactant mixture (15) to produce a first product stream (17) comprising carbonmonoxide, ethylene, and hydrogen; wherein the OCM reaction is performedin the absence of a catalyst, at a temperature ranging from 750°C to 1200°C and a pressure ranging from 0.1 to 5.0 MPa; -providing a second reactant mixture (19) comprising hydrocarbons containingtwo or more carbon atoms in their molecular structure; -performing a steam cracking reaction on the second reactant mixture (19) inthe absence of a catalyst to produce a second product stream (21) comprisingethylene and hydrogen; -separating at least a part of the carbon monoxide contained in, or originatingfrom, the first product stream (17) to produce a CO-rich stream (71); and- performing a water-gas shift reaction on the CO-rich stream to produce an H2-rich stream (83);wherein the OCM reaction and the steam cracking reaction are performed simultaneously in separate sections of a single reactor (1) having two reactorsections (3, 5), the reactor (1) being configured to transfer the heat generated bythe OCM reaction in the first reactor section (3) to the second reactor section (5)for the steam cracking reaction.

2. The process according to claim 1 characterized in that the reactor (1) v is a shell-tubereactor comprising a first reactor section (3) being a shell and a second reactorsection (5) comprising one or more inner tubes; and in that the OCM reaction takesplace in the shell and the steam cracking reaction takes place in the one or more inner tubes.

3. The process according to claim 2 is characterized in that the shell-tube reactor is aco-current reactor comprising a first reactor section (3) being a shell and a second reactor section (5) comprising one or more inner tubes placed inside the shell.

4. The process according to claim 3 is characterized in that the first reactant mixture(15) and / or the second reactant mixture (19) is pre-heated before entering the reactor; with preference, to a temperature ranging from 350 to 700°C.

5. The process according to any one of claims 1 to 4 is characterized in that the secondproduct stream (21) comprises methane and in that methane is separated from thesecond product stream (21) and recycled to the first reactant mixture (15).

6. The process according to any one of claims 1 to 5 is characterized in that the steamcracking reaction is performed on a second reactant mixture (19) comprising one or more selected from ethane, propane, butane, liquefied petroleum gas, naphtha, and diesel; with preference, the second reactant mixture (19) comprises ethane and / or propane.

7. The process according to any one of claims 1 to 6 is characterized in that the steamcracking reaction is performed at a temperature ranging from 750 to 900°C and / or ata pressure ranging from 0.1 to 0.5 MPa.

8. The process according to any one of claims 1 to 7 is characterized in that the firstreactant mixture (15) shows an operating molar ratio CH4:O2 ranging from 2.5 to 20.0, preferably, 2.5 to 10.0.

9. The process according to any one of claims 1 to 8 is characterized in that it comprisestransferring a part of the heat generated by the water gas-shift reaction to pre-heatthe first reactant mixture (15) and / or the second reactant mixture (19) before entering the reactor (1).

10. The process according to any one of claims 1 to 9 is characterized in that the pressureof the first product stream (17) is higher than the pressure of the second productstream (21) and in that the process comprises a compression step wherein thesecond product stream (21) is compressed to at least the pressure of the first product stream (17) and subsequently mixed with the first product stream (17) to form a mixed stream (25) before the step of separating at least a part of the carbon monoxide contained in, or originating from, the first product stream (17) to produce a CO-rich stream (71).

11. The process according to any one of claims 1 to 10 is characterized in that, beforethe step of separating at least a part of the carbon monoxide contained in, ororiginating from, the first product stream (17) to produce a CO-rich stream; thesecond product stream (21) is submitted to a purification step to remove acid gas andto form a purified stream (41); with preference the purification step comprises a sub- step of amine washing followed by a sub-step of caustic washing.

12. The process according to any one of claims 11 is characterized in that the first productstream (17) and the purified stream (41) are mixed to form a mixed stream (25) before the step of separating at least a part of the carbon monoxide contained in, or originating from, the first product stream (17) to produce a CO-rich stream.

13. The process according to any one of claims 1 to 12 is characterized in that, beforethe step of separating at least a part of the carbon monoxide contained in or originating from the first product stream (17) to produce a CO-rich stream; at least one stream selected from the first product stream (17), the second product stream (21) and a mixed stream (25) is submitted to a step of hydrocarbon recovery torecover a stream containing C2+ products; with preference, the step of hydrocarbonrecovery is performed by cryogenic separation.

14. The process according to any one of claims 1 to 13 is characterized in that the stepof separating at least a part of the carbon monoxide contained in the first product stream (17) to produce a CO-rich stream also produces a hydrogen-rich stream.

15. The Process according to any one of claims 1 to 14 is characterized in that the stepof separating at least a part of the carbon monoxide contained in, or originating from,the first product stream (17) to produce a CO-rich stream is performed by cryogenic distillation, or by pressure swing adsorption (PSA) or by COSORB process; preferably by PSA adsorption.

16. An installation to produce the process according to any one of claims 1 to 15 ischaracterized in that it comprises, fluidically connected in the following order:- a reactor (1) having two reactor sections (3, 5) and configured to transfer theheat generated in a reactor section to another reactor section; -an optional compression unit (27)- an optional purification unit (35)- an optional drying unit (43)- a hydrocarbon separation unit (51)- a CO removal unit (59);- a water-gas shift unit (73); and- an optional hydrogen separation unit (83)with preference, the reactor is a shell-tube reactor comprising a first reactorsection (3) being a shell, and a second reactor section (5) comprising one or moreinner tubes.

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